Continous process for the production of n-methyl-diethanolamine (MDEOA)

Abstract

Process for continuously preparing N-methyl-diethanolamine (MDEOA), the process comprising: a) feeding diethanolamine (DEOA), a methanolic formaldehyde solution and hydrogen into a reactor equipped with a heterogenous hydrogenation catalyst; and b) subjecting DEOA, hydrogen and formaldehyde to a reductive amination in the presence of the heterogenous hydrogenation catalyst to obtain MDEOA; and c) withdrawing from the reactor a reactor effluent comprising the MDEOA.

Description

[0001] CONTINOUS PROCESS FOR THE PRODUCTION OF N-METHYL-DIETHANOLAMINE (MDEOA)

[0002] FIELD OF THE INVENTION

[0003] The invention relates to a process for continuously preparing N-methyl-diethanolamine (MDEOA).

[0004] BACKGROUND OF THE INVENTION

[0005] MDEOA is for instance used in scrubbing liquids for the removal of acid compounds from gaseous streams, such as natural gas. It is produced on an industrial scale via the reaction of ethylene oxide with methylamine.

[0006] US 5,091,585 (Huntsman) teaches a continuous process for the preparation of a tertiary aliphatic methyl amine comprising continuously passing an alkyl secondary amine, hydrogen and a formaldehyde source over a heterogenous catalyst. The formaldehyde source may be any of a number of components. It may include formaldehyde alone, formalin, Methyl Formcel, and trioxane; and mixtures of these (cl. 3, II. 37-40). There is no explicit teaching of a methanolic formaldehyde solution. The process described therein is particularly suited for producing fatty tertiary alkyl methyl amines from secondary alkyl amines (cl. 3, 1. 3). There is no hint to the production of MDEOA. At least with respect to the two hydroxyl groups and the length of the alkylene or alkyl moieties, respectively, MDEOA is chemically different from such fatty amines.

[0007] DE 26 18 580 (BASF) teaches the continuous production of MDEOA by reacting DEOA with formaldehyde in the presence of hydrogen and a heterogenous hydrogenation catalyst at a reaction pressure of 150 bar and a reaction temperature of 60 to 100 °C using a 30 wt.-% aqueous formaldehyde solution. It further teaches a preferred pressure in the range from 100 to 200 bar and a temperature not exceeding 110 °C (i.e. a temperature in the range of from 40 to 110°C) (see page 8, lines 2 to 4). There is no hint pointing towards a methanolic formaldehyde solution.

[0008] Diethanolamine (DEOA) is obtained on an industrial scale from the reaction of ethylene oxide and ammonia. The resulting product distribution (monoethanolamine, diethanolamine and triethanolamine) particularly depends on the molar ratio of ethylene oxide and ammonia. Nonetheless, on an industrial scale, diethanolamine will always be produced in non-negligible quantities. Due to various reason oversupply of DEOA is expected in the coming years. Therefore, DEOA is considered a promising starting material for MDEOA production in order to establish an economically interesting alternative for MDEOA production via ethylene oxide and methylamine.

[0009] Therefore, it is an object of the present invention to provide an efficient process for the production of N-methyl- diethanolamine (MDEOA) that uses DEOA as a starting material.

[0010] This object has been solved by a process for continuously preparing N-methyl-diethanolamine (MDEOA), the process comprising: a) feeding diethanolamine (DEOA), a methanolic formaldehyde solution and hydrogen into a reactor equipped with a heterogenous hydrogenation catalyst; and b) subjecting DEOA, hydrogen and formaldehyde to a reductive amination in the presence of the heterogenous hydrogenation catalyst to obtain MDEOA; and c) withdrawing from the reactor a reactor effluent comprising the MDEOA.

[0011] Using methanolic instead of aqueous formaldehyde increases MDEOA yield and produces a raw material with low acid number even at increased reaction temperatures.

[0012] DETAILED DESCRIPTION OF THE INVENTION

[0013] A key feature of the present invention is the use of a methanolic formaldehyde solution. The term "methanolic formaldehyde solution” means that formaldehyde is employed in a solution that comprises methanol and preferably also water.

[0014] In such solution, the formaldehyde may exist in different forms. For instance, a certain amount, in general the major part, of formaldehyde exists as a hemiacetal (resulting from formaldehyde and methanol also referred to as hemiformal or 1 -methoxy-methanol) or polyoxymethylene having the formula HO-[CH2O]n-CH3 with n being an integer typically in the range from 2 to 10. The weight percentages (“wt.-%”) specified herein refer to the "theoretical” amount of formaldehyde, methanol, and water, thus, neglecting any potential reactions among formaldehyde, methanol, and water.

[0015] Preferably, the methanolic formaldehyde solution comprises formaldehyde in the range of 20 to 70 wt.-%, methanol in the range of 20 to 50 wt.-% and water in the range of 5 to 45 wt.-%. The wt.-% is based on the total mass of the methanolic formaldehyde solution.

[0016] More preferably, the methanolic formaldehyde solution comprises formaldehyde in the range of 25 to 65 wt.-%, methanol in the range of 25 to 45 wt.-% and water in the range of 5 to 40 wt.-%.

[0017] Even more preferably, the methanolic formaldehyde solution comprises formaldehyde in the range of 30 to 60 wt.-%, methanol in the range of 30 to 45 wt.-% and water in the range of 5 to 35 wt.-%.

[0018] In a preferred embodiment, the amount of formaldehyde, methanol, and water in the methanolic formaldehyde solution is > 90 wt.-%, preferably > 95 wt.-%, more preferably > 98 wt.-%, even more preferably > 99 wt.-%, particularly preferably > 99.5 wt.-%.

[0019] Using a methanolic instead of aqueous solution has the additional advantage that removal of methanol from the reaction mixture by distillation requires less energy than removal of water, which is due to the lower boiling point of methanol. Moreover, using a solution comprising > 30 wt.-% formaldehyde further reduces energy consumption because less solvent is to be removed. The reaction can be carried out adiabatically, isothermally or quasi isothermally (i.e. isoperiboli-cally). Preferably the reaction is carried out with an isoperibolic temperature profile to control the temperature of the reaction within borders of ±15 K, particularly preferably ±10 K.

[0020] The reaction can be carried out in a broad temperature range, for example in the range from 50 to 200 °C (such as 60 to 200 °C, 80 to 200 °C, 90 to 200 °C or even 100 to 200 °C). It has been found that an increase of the temperature to above 110 °C (the upper limit as taught in DE 26 18 580) results in an increased MDEOA yield. Thus, preferably the reaction temperature is in the range from > 110°C to 200°C. More preferably the reaction temperature is in the range from 115 to 190°C, even more preferably 120 to 180 °C. A reaction temperature of 200 °C is usually not exceeded due to the risk of uncontrolled decomposition reactions in case of a breakdown of the reactor cooling system.

[0021] The above specified ranges are to be understood that the temperature in the reactor is in the respective ranges for either operation mode (i.e. adiabatically, isothermally or quasi isothermally).

[0022] It has been observed that operating the reaction at a temperature above 110 °C results in increased acid formation. In this temperature range using methanolic instead of aqueous formaldehyde solution provides the additional advantage to reduce acid formation and thus helps to keep the amount of acids at an acceptable level. Acid components for example accelerate the deterioration of certain parts of the reactor (e.g. corrosion) and the catalyst. Therefore, their reduction is of crucial importance for the technical efficiency of the process.

[0023] The reaction can be carried out in a broad pressure range, for example a pressure in the range from 60 to 300 bara, preferably 70 to 250 bara, more preferably 80 to 230 bara, even more preferably 90 to 200 bara or even 100 to 190 The pressure specified herein refers to the absolute pressure given in the unit "bara”. The unit "bar” and can be converted to "Pa”, "hPa” or "MPa” (1 bar = 100,000 Pa = 1000 hPa = 0.1 MPa).

[0024] The reaction according to the present invention is carried out in the presence of hydrogen (i.e. molecular hydrogen). The hydrogen is fed continuously to the reactor. The amount of hydrogen fed to the reaction is usually in the range from 50 to 1400 NL / (Lcat. ■ h), preferably 60 to 1200, more preferably 70 to 1000 NL / (Lcat. ■ h), even more preferably 80 to 800 NL / (Lcat. ■ h), particularly preferably 90 to 600 NL / (Lcat. ■ h), more particularly preferably 100 to 400 NL / (Lcat. ■ h).

[0025] The amount of hydrogen is referred to the volume (bed volume) of the heterogenous hydrogenation catalyst in liter (abbreviated as Lcat.) and hour (h).

[0026] NL means standard liters, i.e. liter (of molecular hydrogen) under standard conditions (S.T.P). Standard conditions being understood as follows:

[0027] Standard pressure = 101 325 Pa = 1 013,25 hPa = 101 ,325 kPa = 1 ,01325 bar. Standard temperature = 273.15 K = 0 °C.

[0028] The molar conversion of DEOA is usually in the range from 70 to 100%, preferably 90 to 100%, preferably 95 to 100%, even more preferably 99 to 100%. The conversion refers to the molar amount of DEOA being consumed in the reaction.

[0029] The reaction is usually carried out in a reactor or a plurality of reactors. Preferably the reaction is carried out in a tube reactor, particularly a tube-bundle reactor or a single-stream plant. In case more than one reactor is employed, an interconnection in series or in parallel is possible. In case of a single-stream plant the tube reactor in which the reaction is carried out preferably consists of a plurality of (e.g. two or three) individual tube reactors connected in series. An example for a parallel interconnection is a tube bundle reactor. In principle, it is also possible to combine both kinds of interconnection. For example, by operating two or more tube bundle reactors in series.

[0030] The DEOA can be employed undiluted or diluted in a solvent. Examples of suitable solvents are water, alcohols, and ethers. A preferred solvent is selected from water, methanol, and tetrahydrofuran. Preferably the solvent is selected from water and methanol. Mixtures of respective solvents are also possible (for example a mixture of water and methanol). In case a diluted DEOA is employed, the amount of DEOA is usually in the range of 1 to 99 wt.-%, preferably 40 to 95 wt.-%, more preferably 50 to 90 wt.-%, even more preferably 60 to 89 wt.-% or even 70 to 88 wt.- %, based on the diluted DEOA.

[0031] The molar ratio of formaldehyde to DEOA in the feed to the reactor is usually in the range from 0.8:1 to 2:1, preferably 0.9:1 to 1.5:1, more preferably 1.05:1 to 1.4:1. It being understood, that the molar ratio refers to the molecules (not their functional groups) prior to the reaction. The molar ratio refers to fresh DEOA and fresh formaldehyde. Any DEOA and formaldehyde that might be contained in a recycle stream to the reactor is not considered in the determination of such molar ratio. It is possible to add formaldehyde and DEOA at a reactor input and / or to add them at various inlets being distributed over the length of a reactor between its input and output. In case of an interconnection in series, formaldehyde and or DETA may also be added between two reactors. In case of more than one DEOA feed and / or formaldehyde feed to the reactor, the molar ratio is calculated from the sum over all molar feed rates of formaldehyde and the sum over all molar feed rates of DEOA. Formaldehyde and DEOA can be fed separately to the reactor, i.e. via separate inputs. It is also possible to mix them before feeding them into the reactor.

[0032] The reaction is carried out continuously. "Continuously” means a process in which the starting materials (i.e. diethanolamine (DEOA), methanolic formaldehyde solution and hydrogen) are being processed continuously for a period of time (for example 10 days), which is in contrast to a batch or semi-batch process where the materials are processed in stages. Preferably, the reaction is carried out in the liquid phase, which means that the starting materials are in the liquid phase (DEOA and methanolic formaldehyde solution). Most of the hydrogen is in the gaseous phase and only a small amount in the liquid phase.

[0033] The liquid hourly space velocity (LHSV) over the heterogenous hydrogenation catalyst is typically in the range from 0.2 to 1.0 kg of DEOA / ( at h), preferably 0.2 to 0.95 kg of DEOA / (Lcat h), more preferably 0.2 to 0.8 kg of DEOA / (Lcat h). The LHSV refers to the amount of DEOA in kg which is fed over the volume (bed volume) of the heterogenous hydrogenation catalyst in liter (abbreviated as Lcat.) per hour (h). The LHSV refers to the fresh DEOA. Any DEOA that might be contained in a recycle stream to the reactor is not considered in the determination of the LHSV.

[0034] Besides MDEOA the reactor effluent as per step c) of the present invention also comprises hydrogen, methanol, and water. The water is formed during the reaction and may further be comprised in the methanolic formaldehyde solution.

[0035] In one embodiment the process according to the invention is operated in a recycle gas mode. This means that low volatile components, in particular hydrogen are separated from the reactor effluent and recycled to the reaction (i.e. the reactor). Such removal of volatile components is hereinafter referred to as "degassing”. One obtains a liquid stream hereinafter referred to as "degassed stream”, comprising the MDEOA and a gaseous stream, comprising the hydrogen, hereinafter referred to as "recycle gas stream”. The separation is usually carried out in a high-pressure separator into which the reactor effluent is fed. A high-pressure separator is a device (for instance a vessel) that is operated at a pressure slightly below the reaction pressure. Usually, the pressure in the high-pressure separator is 0.1 to 10 bar, preferably 0.2 to 5 bar below the reaction pressure. The recycle gas stream comprises preferably at least 10% by volume, particularly from 50 to 100% by volume, very particularly from 80 to 100% by volume of H2. The flow rate of the recycle gas stream is preferably in the range from 40 to 1500 m3 (at operating pressure) / [m3 of catalyst (bed volume)' h], in particular in the range from 100 to 700 m3 (at operating pressure) / [m3 of catalyst (bed volume)'h]. To avoid the accumulation of any contaminants comprised in the recycle gas, the recycle gas stream is only partially recycled to the reactor. A small amount, usually 0.1 to 10 wt.-% of the total recycle gas stream obtained from the high-pressure separator is purged. The purged recycle gas is for instance subjected to incineration.

[0036] Preferably the reactor effluent is further subjected to step d), e) and optionally f) as follows: d) subjecting the reactor effluent to degassing to obtain a degassed stream and a recycle gas stream comprising hydrogen; e) partially recycling the recycle gas stream to the reactor; and f) optionally partially recycling the degassed stream to the reactor.

[0037] As outlined above, the degassing of step d) particularly serves the purpose of removing and recycling hydrogen.

[0038] Thus, with respect to its composition, the degassed stream as per step f) corresponds to the reactor effluent which is depleted in hydrogen. Preferably step f) is carried out. Any such recycling is only partial as otherwise no product stream remains. The recycle of the degassed stream helps to control the temperature increase over the length of the reactor (i.e. lower rise of the temperature). Surprisingly it also leads to higher MDEOA selectivity.

[0039] Both, the recycling according to step e) as well as to optional step f) are usually to the inlet of the reactor.

[0040] Preferably step f) is carried out and the weight ratio of the degassed stream to the combined amount of DEOA and methanolic formaldehyde solution being fed to the one or more reactor(s) is in the range of 1 :1 to 10:1, more preferably 2:1 to 8:1 . Such weight ratio is calculated based on the mass flow rate of the degassed stream as per step d), and the respective DEOA and methanolic formaldehyde solution streams being fed to the reactor in accordance with step a). It is to be noted that such streams may also comprise other components. For example, the methanolic formaldehyde solution stream may comprise water (as further specified above) which also contributes to the mass flow rate. In the same way the degassed stream may also comprise water, methanol and certain by-products, which contribute to the mass flow rate. The same is also true with respect to any impurities, which may be comprised in the fresh DEOA.

[0041] It is to be noted, that any starting material that may be comprised in the recycle to the reactor does not contribute to the calculation of the molar ratio of formaldehyde to DEOA, because such ratio is based on fresh formaldehyde and DEOA only.

[0042] In one embodiment, the process according to the present invention comprises steps g) and h) as follows: g) subjecting the degassed stream to distillation to remove low-boilers and to obtain a raw MDEOA stream; and h) subjecting the raw MDEOA stream to distillation to remove high-boilers and to obtain a product MDEOA stream.

[0043] Either step is usually conducted in a distillation column applying appropriate distillation conditions (such as pressure, temperature, and reflux ratio) to provide for sufficient separation efficiency to remove low- and high-boilers, respectively. The person skilled in the art can easily select a suitable distillation column having an appropriate number of theoretical steps.

[0044] A low-boiler is any component that has a lower boiling point than MDEOA at the applied distillation pressure. Low boilers usually comprise water, methanol, and small amounts of unreacted formaldehyde and / or DEOA. A high-boiler is any component that has a higher boiling point than MDEOA at the applied distillation pressure. Without wanting to be bound by any theory, high boilers are believed to comprise dimers and oligomers of DEOA that form as unwanted by-products.

[0045] The product MDEOA stream usually has a purity of > 95 wt.-%, preferably >98 wt.-%, more preferably >99 wt.-% based on the total weight of the product MDEOA stream. The process according to the invention is conducted in the presence of a heterogeneous hydrogenation catalyst. The term heterogeneous catalyst designates a solid catalyst, preferably in the form of particles, which is brought into contact with the reaction mixture comprising the starting materials, any intermediates and MDEOA already obtained.

[0046] Unless explicitly provided otherwise, the heterogenous hydrogenation catalyst is hereinafter also referred to as "catalyst”. Preferred embodiments of the catalyst used in the process according to the present invention are detailed below.

[0047] Any heterogeneous catalyst that has sufficient hydrogenation activity can be used for the manufacture of MDEOA in accordance with the present invention. It may be a supported or an unsupported catalyst. An unsupported catalyst is preferred.

[0048] The heterogeneous hydrogenation catalyst can for instance be prepared by applying precipitation or impregnation methods. Suitable heterogeneous hydrogenation catalysts and respective methods for their production are for instance taught in EP 2043996 B1 , WO 2011 / 067200 A1, WO 2011 / 067199 A1 and EP 2780109 B1 (all BASF).

[0049] Preferably, the heterogeneous hydrogenation catalyst is devoid or substantially devoid of any palladium. The amount of palladium is preferably less than 0.5 wt.-%, preferably less than 0.05 wt.-%, more preferably less than 0.01 wt.-%, based on the total weight of the heterogeneous hydrogenation catalyst.

[0050] The heterogeneous hydrogenation catalyst is installed in the reactor, preferably as fixed bed.

[0051] In a preferred embodiment, the heterogeneous hydrogenation catalyst comprises cobalt. More preferably it comprises cobalt and one or more metal(s) selected from manganese, phosphorus, copper, molybdenum, and an alkali metal, wherein sodium is preferred.

[0052] In another preferred embodiment, the heterogeneous hydrogenation catalyst comprises cobalt, manganese, and phosphorus. More preferably it comprises cobalt, manganese, and phosphorus, and one or more metal (s) selected from copper, molybdenum, and an alkali metal (preferably sodium).

[0053] In another preferred embodiment, the heterogeneous hydrogenation catalyst comprises cobalt, manganese, phosphorus, and an alkali metal, preferably sodium. More preferably it comprises cobalt, manganese, and phosphorus, and one or more metal(s) selected from copper, molybdenum, and an alkali metal (preferably sodium).

[0054] The weight percentages (wt.%) as specified below are based on the total weight of the heterogeneous hydrogenation catalyst. In another preferred embodiment, the heterogeneous hydrogenation catalyst comprises in the range from:

[0055] 5 to 90, preferably 10 to 90, more preferably 15 to 90 wt.-% cobalt, 0 to 30, preferably 0 to 20, more preferably 0 to 10 wt.-% manganese,

[0056] 0 to 30, preferably 0 to 10, more preferably 0 to 5 wt.-% phosphorus,

[0057] 0 to 10, preferably 0 to 8, more preferably 0 to 5 wt.-% of an alkali metal, preferably sodium, 0 to 40, preferably 0 to 30, more preferably 0 to 20 wt.-% copper, and

[0058] 0 to 30, preferably 0 to 20, more preferably 0 to 10 wt.-% molybdenum.

[0059] In another preferred embodiment, the heterogeneous hydrogenation catalyst comprises in the range from:

[0060] 5 to 90, preferably 10 to 90, more preferably 15 to 90 wt.-% cobalt,

[0061] 0.1 to 30, preferably 0.2 to 20, more preferably 1 to 10 wt.-% manganese, 0 .01 to 30, preferably 0.02 to 10, more preferably 0.05 to 5 wt.-% phosphorus, 0 to 10, preferably 0 to 8, more preferably 0 to 5 wt.-% of an alkali metal, preferably sodium, 0 to 40, preferably 0 to 30, more preferably 0 to 20 wt.-% copper, and 0 to 30, preferably 0 to 20, more preferably 0 to 10 wt.-% molybdenum.

[0062] In another preferred embodiment, the heterogeneous hydrogenation catalyst comprises in the range from:

[0063] 5 to 90, preferably 10 to 90, more preferably 15 to 90 wt.-% cobalt,

[0064] 0.1 to 30, preferably 0.2 to 20, more preferably 1 to 10 wt.-% manganese,

[0065] 0.01 to 30, preferably 0.02 to 10, more preferably 0.05 to 5 wt.-% phosphorus,

[0066] 0.01 to 10, preferably 0.02 to 8, more preferably 0.1 to 5 wt.-% of an alkali metal, preferably sodium, 0 to 40, preferably 0 to 30, more preferably 0 to 20 wt.-% copper, and 0 to 30, preferably 0 to 20, more preferably 0 to 10 wt.-% molybdenum.

[0067] The heterogeneous hydrogenation catalyst as specified in any of the three preceding preferred embodiments preferably also contains oxygen. In a preferred embodiment the heterogeneous hydrogenation catalyst comprises oxygen and the cumulated amount of oxygen, cobalt, copper, manganese, molybdenum, phosphorus, and alkali metal is > 80 wt.-%, preferably > 90 wt.-%, more preferably > 95 wt.-% even more preferably > 97 wt.-%, particularly preferably > 98 wt.-%, based on the total weight of the heterogeneous hydrogenation catalyst.

[0068] In another preferred embodiment, the heterogeneous hydrogenation catalyst comprises in the range from:

[0069] 40 to 90, preferably 50 to 90, more preferably 60 to 90 wt.-% cobalt, 0.1 to 30, preferably 0.2 to 20, more preferably 1 to 10 wt.-% manganese, 0 .01 to 30, preferably 0.02 to 10, more preferably 0.05 to 5 wt.-% phosphorus, and 0.01 to 10, preferably 0.02 to 8, more preferably 0.1 to 5 wt.-% of an alkali metal, preferably sodium.

[0070] In another preferred embodiment, the heterogeneous hydrogenation catalyst comprises in the range from:

[0071] 5 to 90 wt.-%, preferably 10 to 60 wt.-% cobalt, 1 to 40 wt.-%, preferably 2 to 30 wt.-% copper,

[0072] 0.1 to 30 wt.-%, preferably 1 to 10 wt.-% manganese, 0.1 to 30 wt.-%, preferably 1 to 10 wt.-% molybdenum, and 0.05 to 30 wt.-%, preferably 0.1 to 5 wt.-% phosphorus.

[0073] In case the heterogeneous hydrogenation catalyst comprises an alkali metal, such alkali metal is preferably selected from lithium, sodium and kalium. Particularly preferred is sodium. In principle, the catalyst can also comprise more than one alkali metal. In case of any embodiment of the catalyst, wherein the alkali metal is preferably selected from lithium, sodium and kalium or particularly preferably comprises sodium, the overall amount of alkali metals does not exceed the upper limit of 10, 8, or 5 wt.-%, respectively.

[0074] The heterogeneous hydrogenation catalyst as specified in the preceding preferred embodiment preferably also contains oxygen. In a preferred embodiment the heterogeneous hydrogenation catalyst comprises oxygen and the cumulated amount of oxygen, cobalt, copper, manganese, molybdenum, phosphorus, and alkali metals is > 80 wt.- %, preferably > 90 wt.-%, more preferably > 95 wt.-% even more preferably > 97 wt.-%, particularly preferably > 98 wt.-%, based on the total weight of the heterogeneous hydrogenation catalyst.

[0075] The preparation of such catalysts is for instance taught in EP0636409B1 (BASF) or DE 2321101 (BASF). Accordingly, the catalyst is obtained by precipitation, followed by calcination. The catalyst thus obtained is activated by reduction in a hydrogen stream. It is to be noted, that any “wt.-%” as specified herein with respect to the composition of the heterogeneous hydrogenation catalyst refers the total weight of the heterogeneous hydrogenation catalyst after the last of any heat treatments (for instance calcination) and prior to its reduction with hydrogen.

[0076] It follows from the nature of the preparation method, that the respective metals at least partially exist in an oxidized from. Nonetheless, the presence of a certain amounts of such metals in elementary form is not excluded. For instance, in the course of the preparation, one could, besides using respective metal nitrites, also apply certain amount(s) of respective metal(s) in elementary form. Usually, more than 90 wt.-%, preferably more than 95 wt.-%, more preferably more than 99 wt.-% (or even more than 99.5 wt.-%) of any respective metal exists in oxidized from. The phosphorus usually exists substantially in an oxidized form. Usually, one would not apply elementary phosphor in the catalyst preparation, but phosphorus in an oxidized form only (in particular phosphoric acid). Moreover, the formation of elementary phosphorus during precipitation or calcination is unlikely to happen. Thus, preferably more than 99 wt.-%, more preferably more than 99.5 wt.-%, even more preferably more than 99.9 wt.-% of the respective phosphorus exist in oxidized form.

[0077] The following examples only serve for the purpose of the illustration of the present invention and shall therefore not limit it in whatsoever kind. EXAMPLES

[0078] The catalyst A was prepared in accordance with Example 1 of EP0636409B1 (BASF). Its composition (based on the total weight of the heterogenous hydrogenation catalyst prior to its reduction with hydrogen) is as follows:

[0079] Elementary metal / phosphorus in g / 100 g catalyst: Co (72.0 g), Mn (3.7 g), P (1.0 g), Na (0.2 g) in 100 g catalyst.

[0080] Catalyst B:

[0081] The catalyst B was prepared in accordance with Example 1 of DE 2321101 (BASF). Its composition (based on the total weight of the heterogenous hydrogenation catalyst prior to its reduction with hydrogen) is as follows:

[0082] Elementary metal / phosphorus in g / 100 g catalyst: Co (16.5 g), Cu (15.2 g), Mn (2.5 g), Mo (2.4 g), P (0.9 g), to MDEOA with different reaction conditions:

[0083] A heated reactor with internal diameter 10 mm and a total volume of 95 mL was charged in the lower section with 2-3 wire mesh rings, followed by heterogeneous and finally 2-3 wire mesh rings. Prior to the reaction, the catalyst (catalyst A) was activated at max. 270° C. Thereby, the catalyst was heated to 250 °C with a hydrogen flow of 90 NL / h. After 12 h at 250 °C, the reactor was cooled to 40 °C. A pressure of 150 bar was employed. The reactor was operated in a recycle gas mode. 25 g / h of diethanolamine (DECA) were metered through the reactor. The reactor was kept at a temperature and absolute pressure as set forth in table 1 below. Additionally, methanolic formaldehyde (55 wt.-% formaldehyde, 35 wt.-% methanol, and 10 wt.-% water) was metered through a second pump. At different times, samples were taken from the reaction mixture and analyzed by means of gas chromatography. For this purpose, an "Volamin” GC column was used with the following parameters: Volamin (30 m x 0,32 mm x 4,5 pm) 80°C -5min-5°C / min-260°C / 20min.

[0084] The results are presented in table 1 below. Table 1 - Results

[0085] * Comparative Example (reaction conditions as per example 2 of DE 26 18580)

[0086] ** Comparative Example (yield optimized reaction conditions using aqueous formaldehyde solution)

[0087] Example 6:

[0088] The same set-up as per Examples 1 to 5 was applied except that (1) the DECA was employed as a solution (75 wt.- % DECA and 25 wt.-% methanol) and (2) a recycle feed was in place which recycled reactor effluent after the high- pressure separator to the reactor inlet, the mass flow ratio of the recycled feed to DECA and methanolic formaldehyde feed (combined flow rates) being 3 : 1. The resulting MDEOA selectivity was 99.1 % and no DECA could be detected in the reactor effluent.

[0089] Example 7:

[0090] The same set-up as per Examples 1 to 5 was applied except that (1) the DECA was employed as a solution (85 wt.- % DECA and 15 wt.-% water), (2) catalyst B was used, (3) a methanolic formaldehyde (weight ratio of formaldehyde, methanol, and water being 1 :1 :1), and (4) a recycle feed was in place which recycled reactor effluent after the high- pressure separator to the reactor inlet, the mass flow ratio of the recycled feed to DEOA and methanolic formaldehyde feed (combined flow rates) being 3 : 1. The resulting MDEOA selectivity was 98.9 %. 0.11% DEOA were detected in the reactor effluent. Discussion of results:

[0091] A comparison of examples 1 and 2 shows that using methanolic formaldehyde instead of aqueous formaldehyde increases MDEOA yield. An additional boost of MDEOA yield is obtained by increasing the reaction temperature (see example 3). Such temperature increase however causes a rise of the acid value in the reaction mixture (from 0.9 and 1.4 to 4.4). It has been found that such acid formation can be repressed when methanolic instead of aqueous formaldehyde solution is used. The acid value in example 5 (9.0 using aqueous formaldehyde) is more than twice as high than the acid value in example 4 (3.8 using methanolic formaldehyde). An additional significant boost of MDEOA yield is obtained when a product recycle (comparison of examples 4, 5 and 6) is realized. Example 7 shows, that one also obtains high MDEOA selectivity, when a different catalyst (i.e. catalyst B) and a methanolic formaldehyde solution having about 33.3 wt.-% formaldehyde is employed.

Claims

CLAIMS1. Process for continuously preparing N-methyl-diethanolamine (MDEOA), the process comprising: a) feeding diethanolamine (DECA), a methanolic formaldehyde solution and hydrogen into a reactor equipped with a heterogenous hydrogenation catalyst; and b) subjecting DECA, hydrogen and formaldehyde to a reductive amination in the presence of the heterogenous hydrogenation catalyst to obtain MDEOA; and c) withdrawing from the reactor a reactor effluent comprising the MDEOA.

2. The process according to claim 1 , wherein the methanolic formaldehyde solution comprises formaldehyde in the range from 20 to 70 wt.-%, methanol in the range from 20 to 50 wt.-% and water in the range from 5 to 45 wt.-%.

3. The process according to any of the preceding claims, wherein the reaction temperature is in the range from 50 to 200 °C (such as 60 to 200 °C, 80 to 200 °C, 90 to 200 °C or even 100 to 200 °C), preferably > 110°C to 200°C, more preferably 115 to 190°C, even more preferably 120 to 180 °C.

4. The process according to any of the preceding claims, wherein the reaction pressure is in the range from 60 to 300 bara, preferably 70 to 250 bara, more preferably 80 to 230 bara, even more preferably 90 to 200 bara or even 100 to 190 bara.

5. The process according to any of the preceding claims, wherein the molar ratio of formaldehyde to DEOA in the feed to the reactor is in the range from 0.8:1 to 2:1, preferably 0.9:1 to 1.5:1, more preferably 1.05:1 to 1.4:1.

6. The process according to any of the preceding claims, wherein the reductive amination is conducted in the liquid phase.

7. The process according to any of the preceding claims, wherein the liquid hourly space velocity (LHSV) over the heterogenous hydrogenation catalyst is in the range from 0.2 to 1 .0 kg of DEOA / (Lkat h), preferably 0.2 to 0.95 kg of DEOA / (Lkat h), more preferably 0.2 to 0.8 kg of DEOA / (Lkat h).

8. The process according to any of the preceding claims, wherein the process comprises steps d),e) and f) as follows: d) subjecting the reactor effluent to degassing to obtain a degassed stream and a recycle gas stream comprising hydrogen; e) partially recycling the recycle gas stream to the reactor; and f) optionally partially recycling a part of the degassed stream to the reactor.

9. The process according to the preceding claim, wherein step f) is carried out and wherein the weight ratio of the degassed stream to the combined amount of DEOA and methanolic formaldehyde solution being fed to the one or more reactor(s) is in the range of 1 :1 to 10:1, more preferably 2:1 to 8:1.

10. The process according to any of the two preceding claims, wherein the process comprises steps g) and h) as follows: g) subjecting the degassed stream to distillation to remove low-boilers and to obtain a raw MDEOA stream; and h) subjecting the raw MDEOA stream to distillation to remove high-boilers and to obtain a product MDEOA stream.11 . The process according to any of the preceding claims, wherein the heterogenous hydrogenation catalyst comprises cobalt.

12. The process according to any of claims 1 to 11, wherein the heterogenous hydrogenation catalyst comprises in the range from:5 to 90, preferably 10 to 90, more preferably 15 to 90 wt.-% cobalt,0 to 30, preferably 0 to 20, more preferably 0 to 10 wt.-% manganese,0 to 30, preferably 0 to 10, more preferably 0 to 5 wt.-% phosphorus,0 to 10, preferably 0 to 8, more preferably 0 to 5 wt.-% of an alkali metal, preferably sodium,0 to 40, preferably 0 to 30, more preferably 0 to 20 wt.-% copper, and0 to 30, preferably 0 to 20, more preferably 0 to 10 wt.-% molybdenum, based on the total weight of the heterogeneous hydrogenation catalyst.

13. The process according to any of claims 1 to 11, wherein the heterogenous hydrogenation catalyst comprises cobalt, manganese, and phosphorus.

14. The process according to any of claims 1 to 11, wherein the heterogenous hydrogenation catalyst comprises cobalt, manganese, phosphorus, and an alkali metal, preferably sodium.

15. The process according to any of claims 1 to 11, wherein the heterogeneous hydrogenation catalyst comprises in the range from:5 to 90, preferably 10 to 90, more preferably 15 to 90 wt.-% cobalt,0.1 to 30, preferably 0.2 to 20, more preferably 1 to 10 wt.-% manganese,0 .01 to 30, preferably 0.02 to 10, more preferably 0.05 to 5 wt.-% phosphorus,0 to 10, preferably 0 to 8, more preferably 0 to 5 wt.-% of an alkali metal, preferably sodium, 0 to 40, preferably 0 to 30, more preferably 0 to 20 wt.-% copper, and0 to 30, preferably 0 to 20, more preferably 0 to 10 wt.-% molybdenum,based on the total weight of the heterogeneous hydrogenation catalyst.